Method for the production of tetrahydrofuran

ABSTRACT

Unsubstituted or alkyl-substituted THF is obtained by catalytic hydrogenation in the gas phase of C 4 -dicarboxylic acids and/or their derivatives using a catalyst comprising &lt;80% by weight, preferably &lt;70% by weight, in particular from 10 to 65% by weight, of CuO and &gt;20% by weight, preferably &gt;30% by weight, in particular from 35 to 90% by weight, of an oxidic support having acid centers, at a hot spot temperature of from 240 to 310° C., preferably from 240 to 280° C., and a WHSV over the catalyst of from 0.01 to 1.0, preferably from 0.02 to 1, in particular from 0.05 to 0.5, kg of starting material/l of catalyst x hour.

The present invention relates to a process for preparing unsubstitutedor alkyl-substituted tetrahydrofuran by catalytic hydrogenation in thegas phase of substrates selected from the group consisting ofderivatives of maleic acid and succinic acid and these acids themselves.For the purposes of the present invention, derivatives are anhydrideswhich, like the acids, may bear one or more alkyl substituents.

The preparation of tetrahydrofuran (THF) by gas-phase hydrogenation ofmaleic anhydride (MA) is a reaction which has been known for many years.Numerous catalyst systems for carrying out this catalytic reaction aredescribed in the literature. Depending on the composition of thecatalyst and the reaction parameters chosen, different productdistributions are obtained using such catalysts.

It has been able to be shown that the hydrogenation of MA to THF formsfirstly succinic anhydride (SA) and subsequently γ-butyrolactone (GBL)as intermediates and the latter can be hydrogenated further to form1,4-butanediol (BDO). All the abovementioned intermediates canthemselves be used as starting materials for preparing THF.

If GBL and THF bearing alkyl substituents are to be prepared, thealkyl-substituted species corresponding to the abovementioned startingmaterials can be used.

The catalysts used in the hydrogenation frequently comprise chromium,particularly in older processes. This is reflected in the patentliterature in which there are a large number of patents and patentapplications which disclose chromium-containing catalysts for theabove-described hydrogenation reaction, with the hydrogenation in mostcases being restricted to MA as starting material. U.S. Pat. No.3,065,243 discloses a process in which copper chromite is employed ascatalyst. According to the description and examples, this reactionresults in formation of considerable amounts of SA which has to becirculated. As is known, process engineering problems due tocrystallization of SA or succinic acid formed therefrom with subsequentblocking of pipes frequently occur.

Further copper chromite catalysts for the hydrogenation of MA aredisclosed, for example, in U.S. Pat. Nos. 3,580,930, 4,006,165, EP-A 638565 and WO 99/38856.

The catalysts used in U.S. Pat. No. 5,072,009 have the formulaCu_(l)Zn_(b)Al_(c)M_(d)O_(x) in which M is at least one element selectedfrom the group consisting of the elements of groups IIA and IIIA, VA,VIII, Ag, Au, the elements of groups IIIB to VIIB and the lanthanidesand actinides of the Periodic Table of the Elements, b is from 0.001 to500, c is from 0.001 to 500 and d is from 0 to <200 and x corresponds tothe number of oxygen atoms necessary according to valence criteria. Inthe examples, where only chromium-containing catalysts are used, thehydrogenation of MA using the catalysts of that invention forms THF inyields of over 90%.

EP-A 0 404 408 discloses a catalyst whose catalytically active materialcorresponds essentially to the material disclosed in the above-citedU.S. Pat. No. 5,072,009 for hydrogenation of MA. The catalyticallyactive material is used in immobilized form on a support as coatedcatalyst and not as all-active catalyst. In contrast to the materialpresent as all-active catalyst, mainly GBL is formed according to theexamples reported, in which once again only chromium-containingcatalysts are used.

A two-stage catalyst system for the hydrogenation of MA is described inU.S. Pat. No. 5,149,836. The catalyst for the first stage ischromium-free while the catalyst for the second stage is based onCu—Zn—Cr oxides.

Owing to the toxicity of chromium, more modem technologies areincreasing moving away from the use of chromium-containing catalysts.Examples of chromium-free catalyst systems may be found in WO 99/35139(Cu—Zn oxide), WO 95/22539 (Cu—Zn—Zr) and U.S. Pat. No. 5,122,495(Cu—Zn—Al oxide).

A catalyst made up exclusively of copper oxide and aluminum oxide forthe gas-phase hydrogenation of MA to form GBL is disclosed in WO97/24346. This catalyst comprises from 50 to 95% by weight, preferablyfrom 80 to 90% by weight, of copper oxide, from 3 to 30% by weight,preferably from 5 to 15% by weight, of aluminum oxide and optionally abinder. Selectivities to GBL of up to about 98% are achieved in thehydrogenation of MA using such a catalyst.

WO 91/16132 discloses a chromium-free catalyst for the hydrogenation ofMA. The catalyst comprises from 30 to 65% by weight of CuO, from 18 to50% by weight of ZnO and from 8 to 22% by weight of Al₂O₃. Before use inthe hydrogenation reaction, the catalyst is reduced in a hydrogenatmosphere and subsequently activated in a hydrogen atmosphere at notless than 400° C. for at least 8 hours. Such a catalyst gives GBLselectivities of about 90%.

In Catalysis Today 27 (1996), pp. 181 to 186, Castiglioni et al.disclose CuO/ZnO/Al₂O₃ catalysts which give mainly GBL in thehydrogenation of MA; a maximum THF selectivity of 17% is observed.

The use of a catalyst having a similar composition as in WO 97/24346 isalso disclosed in JP 2 233 631. The object of that invention is to carryout the hydrogenation of MA in such a way that THF and BDO are formed asmain products and only small amounts, if any, of GBL, are formed. Thisis achieved by use of catalysts based on mixed Cu—Al oxides and byadherence to particular reaction conditions. General indications ofamounts of the Cu—Al oxide are not given; the examples disclose twocatalyst compositions, one comprising about 46% by weight of CuO and 33%by weight of Al₂O₃ and the other comprising about 36% by weight of CuOand 47.% by weight of Al₂O₃. Use of this catalyst is said to give a THFselectivity of up to 99%, but only when an excess of GBL is employed assolvent. If the hydrogenation is carried out using the same catalyst inthe absence of GBL, the selectivity drops to 76%. According to theexamples, the hydrogenation is carried out at from about 210 to 230° C.and GHSVs of from about 3200 to 9600. The hydrogen/MA ratios are atvalues which are rather unfavorably high for industrial processes,namely from 200 to 800 in the examples.

The hydrogenation of MA under conditions corresponding to those of JP 2233 631 but using a different catalyst is disclosed in JP 2 639 463. Theuse of the catalyst is said to make it possible to prepare BDO and THFby hydrogenation of MA. Use is made here of a copper oxide/zincoxide/aluminum oxide catalyst whose composition is not disclosedquantitatively in the description. The catalysts used according to theexamples have a composition of 20% by weight of CuO, 43.6% by weight ofZnO and 18.1% by weight of Al₂O₃, 32.6% by weight of CuO, 38.1% byweight of ZnO and 9.5% by weight of Al₂O₃, 24.2% by weight of CuO, 36.4%by weight of ZnO and 17.2% by weight of Al₂O₃, 26.4% by weight of CuO,52.9% by weight of ZnO, 7.6% by weight of Al₂O₃ and 1.4% by weight ofCaO or 22.9% by weight of CuO, 44.8% by weight of ZnO and 16.3% byweight of Al₂O₃. The hydrogenation is generally carried out using asolvent such as GBL or dioxane, giving a maximum THF selectivity of 94%.When the reaction is carried out without a solvent, the THF selectivityis no more than 83%.

The technologies on which the above-cited publications are based useprepurified MA which has, after its preparation, generally been freed ofimpurities by distillation as starting material for the hydrogenationreactions. MA is prepared by partial oxidation of particularhydrocarbons, namely benzene, butene mixtures and n-butane, withpreference being given to using the latter. The crude product of theoxidation comprises not only the desired MA but also, in particular,by-products such as water, carbon monoxide, carbon dioxide, unreactedstarting hydrocarbons and acetic and acrylic acids, with theseby-products being independent of the hydrocarbons used in the oxidation.The by-products are usually separated off by complicated methods, forexample by distillation as mentioned above. The purification isnecessary because, in particular, the catalysts used in thehydrogenation process are generally sensitive to such impurities.Deactivation of the catalysts is a problem even when using purified MA,since coating of the catalyst with polymerization products of MA makesit necessary to regenerate the catalyst at generally relatively shortintervals, often about 100 hours. The tendency for deactivation to occuris increased further when polymerizable impurities such as acrylic acidare present. This fact is known to those skilled in the art and is alsodescribed, for example, in the patent applications EP-A 322 140, WO91/16132 and DE-A 240 44 93.

Up to now, the prior art contains only one publication which disclosesthe hydrogenation of MA which has been only roughly prepurified. WO97/43234 discloses the absorption of maleic anhydride from gas streamswhich comprise maleic anhydride and originate from the oxidation ofhydrocarbons by means of absorption media which have a boiling point atleast 30° C. higher, stripping the maleic anhydride from theseabsorption media by means of hydrogen and hydrogenating the hydrogenstream comprising maleic anhydride in the gas phase over a heterogeneouscatalyst. This gives mainly BDO together with small amounts of GBL andTHF. The hydrogenation is carried out in the gas phase at from about150° C. to 300° C. and a pressure of from 5 bar to 100 bar. Catalystsused are promoted copper catalysts as are described in Journal ofCatalysis 150, pages 177 to 185 (1994). These are chromium-containingcatalysts of the Cu/Mn/Ba/Cr and Cu/Zn/Mg/Cr types. Thus, according tothe disclosure of that application, chromium-containing catalysts areused for the hydrogenation of MA grades containing the above-describedimpurities. However, owing to the toxicity of chromium, the use ofchromium-containing catalysts is nowadays avoided as much as possible.In addition, the process forms not only the main product BDO but alsoappreciable amounts of undesired products, namely THF and GBL whicheither have to be separated off and purified or have to be returned tothe hydrogenation. This is frequently associated with undesirable costs,particularly in industrial processes.

Which of the products GBL, THF and BDO obtainable in the hydrogenationof MA is the desired product often depends on the market growth of theseproducts or their downstream products or on the wider product range ofthe respective producer. It can therefore be said that some producersuse the hydrogenation of MA to produce THF while others use it forobtaining GBL and/or BDO.

It is an object of the present invention to provide a process by meansof which THF can be prepared by hydrogenation of MA. This process shouldbe able to be operated continuously using chromium-free catalysts andshould give the largest possible amount of the desired product THF so asto achieve the best possible economics. Furthermore, the catalyst shouldbe able to be employed with MA which has not been laboriouslyprepurified, for example by distillation, and nevertheless have a highstability, i.e. not require frequent regeneration.

We have found that this object is achieved by a process for preparingunsubstituted or alkyl-substituted THF by catalytic hydrogenation in thegas phase of C₄dicarboxylic acids and/or their derivatives using acatalyst comprising <80% by weight, preferably <70% by weight, inparticular from 10 to 65% by weight, of CuO and >20% by weight,preferably >30% by weight, in particular from 35 to 90% by weight, of anoxidic support having acid centers, at a hot spot temperature of from240 to 310° C., preferably from 240 to 280° C., and a WHSV over thecatalyst of from 0.01 to 1.0, preferably from 0.02 to 1, in particularfrom 0.05 to 0.5, kg of starting material/l of catalyst x hour.

For the purposes of the present invention, the term C₄-dicarboxylicacids and their derivatives refers to maleic acid and succinic acidwhich may be unsubstituted or bear one or more C₁-C₆ alkyl substituentsand also the anhydrides of these unsubstituted or alkyl-substitutedacids. An example of such an acid is citraconic acid. Preference isgiven to using the anhydride of a given acid. Particular preference isgiven to using MA as starting material.

The process of the present invention can be operated inexpensively andgives high THF yields and selectivities. This is achieved by adherenceto particular conditions and parameters. It is also possible to carryout the process continuously, which is preferred according to thepresent invention.

An important aspect is the choice of the catalyst, which comprisescopper oxide as main catalytically active constituent. This is appliedto an oxidic support which has to have a suitable number of acidcenters. The amount of oxidic support required depends on the number ofacid centers present therein. A suitable support material having asufficient number of acid centers is aluminum oxide, whose use ispreferred according to an embodiment of the present invention. Accordingto another embodiment of the present invention, preference is given tousing a combination of aluminum oxide with zinc oxide in a weight ratioof from 20:1 to 1:20, preferably from 5:1 to 1:5, as acid supportmaterial. In the case of materials having a large number of such acidcenters, the lower limit of the amount of support consisting of such amaterial is 20% by weight. The amount of copper oxide is <80% by weight.Preferred catalyst compositions comprise <70% by weight of copper oxideand >30% by weight of support; particularly preferred catalysts comprisefrom 10 to 65% by weight of copper oxide and from 35 to 90% by weight ofsupport.

Low copper oxide contents are preferred because of the cost advantageachieved in this way. High yields can be achieved as a result of theacid support materials.

The catalysts used according to the present invention, which areCr-free, may optionally further comprise one or more additional metalsor compounds thereof, preferably oxides, from groups 1 to 14 (IA toVIIIA and IB to IVB according to the old IUPAC nomenclature) of thePeriodic Table of the Elements. If such a further oxide is used,preference is given to using TiO₂, ZrO₂, SiO₂ and/or MgO.

In addition, the catalysts used may contain from 0 to 10% by weight ofan auxiliary. For the purposes of the present invention, auxiliaries areorganic and inorganic materials which contribute to improvedprocessability during catalyst production and/or to an increase in themechanical strength of the shaped catalyst bodies. Such auxiliaries areknown to those skilled in the art; examples include graphite, stearicacid, silica gel and copper powder.

The catalysts can be produced by methods known to those skilled in theart. Preference is given to methods in which the copper oxide isobtained in finely divided form and intimately mixed with the otherconstituents; particular preference is given to precipitation reactions.In such methods, precursor compounds dissolved in a solvent areprecipitated by means of a precipitant in the presence of furthersoluble metal compounds or metal compounds suspended in the solvent,filtered, washed, dried and optionally calcined.

These starting materials can be processed by known methods to form theshaped bodies, for example by extrusion, tableting or by agglomerationprocesses, with or without addition of auxiliaries.

As an alternative, catalysts for use according to the present inventioncan also be produced, for example, by application of the activecomponent to a support, for example by impregnation or vapor deposition.Furthermore, catalysts to be used according to the present invention canbe obtained by shaping a heterogeneous mixture of active component orprecursor compound thereof with a support component or precursorcompound thereof.

In the hydrogenation according to the present invention, in which notonly MA but also other, above-defined C₄-dicarboxylic acids orderivatives thereof can be used as starting material, the catalyst isemployed in reduced, activated form. Activation is carried out by meansof reducing gases, preferably hydrogen or hydrogen/inert gas mixtures,either before or after installation in the reactor in which the processof the present invention is carried out. If the catalyst has beeninstalled in the reactor in oxidic form, activation can be carried outeither before the plant is started up to carry out the hydrogenationaccording to the present invention or during the process, i.e. in situ.Separate activation before starting up the plant is generally carriedout using reducing gases, preferably hydrogen or hydrogen/inert gasmixtures, at elevated temperatures, preferably from 100 to 300° C. Inthe in situ activation, activation is carried out by contact withhydrogen at elevated temperature while running up the plant.

The catalysts are used as shaped bodies. Examples include extrudates,ribbed. extrudates, other extruded shapes, pellets, rings, spheres andcrushed material. The BET surface area of the copper catalysts in theoxidic state is from 10 to 400 m²/g, preferably from 15 to 200 m²/g, inparticular from 20 to 150 m²/g. The copper surface area (N₂Odecomposition) of the reduced catalyst in the installed state is >0.2m²/g, preferably >1 m²/g, in particular >2 m²/g.

In one variant of the invention, catalysts having a defined porosity areused. These catalysts display, as shaped bodies, a pore volume of ≧0.01ml/g for pore diameters of >50 nm, preferably ≧0.025 ml/g for porediameters of >100 nm and in particular ≧0.05 ml/g for pore diametersof >200 nm. Furthermore, the ratio of macropores having a diameterof >50 nm to the total pore volume for pores having a diameter of >4 nmis >10%, preferably >20%, in particular >30%. The use of these catalystsoften makes it possible to achieve high THF yields and selectivities.The porosities reported were determined by mercury intrusion inaccordance with DIN 66133. The data were evaluated in the pore diameterrange from 4 nm to 300 μm.

The catalysts used according to the present invention generally have asufficient operating life. However, should the activity and/orselectivity of the catalyst drop during operation, it can be restored bymeans of measures known to those skilled in the art. These include,preferably, reductive treatment of the catalyst in a stream of hydrogenat elevated temperature. The reductive treatment may, if appropriate, bepreceded by an oxidative treatment. Here, a gas mixture comprisingmolecular oxygen, for example air, is passed at elevated temperaturethrough the catalyst bed. It is also possible to wash the catalyst witha suitable solvent, for example ethanol, THF or GBL, and subsequently todry it in a stream of gas.

Furthermore, adherence to particular reaction parameters is necessary toachieve the THF selectivities according to the present invention.

An important parameter is adherence to a suitable reaction temperature.This is achieved, firstly, by means of a sufficiently high inlettemperature of the starting materials. This is from >220 to 300° C.,preferably from 235 to 270° C. To obtain an acceptable or high THFselectivity and yield, the reaction has to be carried out so that asuitably high reaction temperature prevails in the catalyst bed in whichthe actual reaction takes place. This temperature, known as the hot spottemperature, is established after entry of the starting materials intothe reactor and is in the range from 240 to 310° C., preferably from 240to 280° C. The process is carried out so that the inlet temperature andthe outlet temperature of the reaction gases are below this hot spottemperature. The hot spot is advantageously located in the first half ofthe reactor, particularly in the case of a shell-and-tube reactor. Thehot spot temperature is preferably from 5 to 15° C., in particular from10 to 15° C., above the inlet temperature. If the hydrogenation iscarried out below the minimum temperatures for the inlet and hot spottemperatures respectively, the amount of GBL increases and the amount ofTHF decreases at the same time when using MA as starting material.Furthermore, deactivation of the catalyst due to coating with succinicacid, funaric acid and/or SA during the course of the hydrogenation isobserved at such temperatures. On the other hand, if MA as startingmaterial is hydrogenated at above the maximum temperatures for the inletand hot spot temperatures respectively, the THF yield and selectivitydrop to unsatisfactory values. In this case, increased formation ofn-butanol and n-butane is observed, i.e. the products of a furtherhydrogenation.

The WHSV over the catalyst in the hydrogenation of the present inventionis in the range from 0.01 to 1.0 kg of starting material/l of catalyst xhour. In the case of a possible but not preferred recirculation ofintermediate formed by incomplete hydrogenation, GBL when using MA asstarting material, the WHSV over the catalyst is the sum of freshstarting material fed in and recirculated intermediate. If the WHSV overthe catalyst is increased beyond the specified range, an increase in theproportion of intermediate in the output from the hydrogenation isgenerally observed. The WHSV over the catalyst is preferably in therange from 0.02 to 1, in particular from 0.05 to 0.5, kg of startingmaterial/l of catalyst x hour. In the case of recirculation, the term“starting material” also includes initially formed hydrogenation productwhich is then further hydrogenated after recirculation to form aproduct, i.e., for example, GBL when MA is used in the hydrogenationreaction.

The hydrogen/starting material molar ratio is likewise a parameter whichhas an important influence on the product distribution and the economicsof the process of the present invention. From an economic point of view,a low hydrogen/starting material ratio is desirable. The lower limit is5, but higher hydrogen/starting material molar ratios of from 20 to 400are generally employed. The use of the above-described catalysts usedaccording to the present invention and adherence to the above-describedtemperatures allows the use of favorable, low hydrogen/starting materialratios which are preferably in the range from 20 to 200, more preferablyfrom 40 to 150. The most favorable range is from 50 to 100.

To set the hydrogen/starting material molar ratios used according to thepresent invention, part, advantageously the major part, of the hydrogenis circulated. For this purpose, circulating gas compressors known tothose skilled in the art are generally used. The amount of hydrogenconsumed chemically by the hydrogenation is replaced. In a preferredembodiment, part of the circulating gas is bled off to remove inerts,for example n-butane. The circulated hydrogen can also be used, ifnecessary after preheating, for vaporizing the starting material stream.

The volume flow of the reaction gases, generally expressed as GHSV (GasHourly Space Velocity) is also an important parameter in the process ofthe present invention. The GHSV in the process of the present inventionis in the range from 100 to 10,000 Standard m³/m³h, preferably from 1000to 3000 Standard m³/m³h, in particular from 1100 to 2500 Standardm³/m³h.

The pressure at which the hydrogenation of the present invention iscarried out is in the range from 1 to 30 bar, preferably from 2 to 9bar, in particular from 3 to 7 bar.

All products which are not condensed or only incompletely condensed oncooling the gas stream leaving the hydrogenation reactor are circulatedtogether with the circulating hydrogen. These are predominantly THF,water and by-products such as methane and butane. The coolingtemperature is from 0 to 60° C., preferably from 20 to 45° C. The THFcontent of the circulating gas is from 0.1 to 5% by volume, inparticular from 1 to 3% by volume.

It is known from the literature that THF and GBL can be hydrogenated bymeans of hydrogen in the presence of copper catalysts to form n-butanol.The process of the present invention is notable for the fact that,despite the high proportions of THF in the circulating gas, which cangenerally be readily hydrogenated further to n-butanol, THF yields ofover 90%, sometimes even over 95%, are achieved.

Possible types of reactor are all apparatuses suitable forheterogeneously catalyzed reactions involving a gaseous startingmaterial and a product stream. Preference is given to tube reactors,shaft reactors or reactors with internal heat removal, for exampleshell-and-tube reactors; it is also possible to use a fluidized bed.Particular preference is given to using shell-and-tube reactors. It ispossible to use a plurality of reactors connected in parallel or inseries. In principle, streams can be fed in between the catalyst beds.Intermediate cooling between or in the catalyst beds is also possible.When using fixed-bed reactors, dilution of the catalyst by inertmaterial is possible.

The gas stream leaving the reactor is cooled to from 10 to 60° C. Thereaction products are condensed out in this way and are passed to aseparator. The uncondensed gas stream is taken off from the separatorand passed to the circulating gas compressor. A small amount ofcirculating gas is bled off. The condensed reaction products are takencontinuously from the system and passed to work-up. By-products presentin the condensed liquid phase are mainly n-butanol together with smallamounts of propanol.

The hydrogenation product is then fractionally distilled to separate theazeotrope of water and unsubstituted or alkyl-substituted THF from anyby-product, for example GBL. The water-containing THF is dewatered in aknown manner and worked up by distillation to give THF which meetsspecifications. By-product such as GBL is returned to the hydrogenationor worked up by distillation.

In the process of the present invention, starting materials of differingpurities can be used in the hydrogenation reaction. Of course, it ispossible to use a high-purity starting material, in particular MA, inthe hydrogenation reaction. However, the catalyst used according to thepresent invention and the other reaction conditions selected accordingto the present invention make it possible to use starting materials, inparticular MA, which is/are contaminated by the customary compoundsformed in the oxidation of benzene, butenes or n-butane and by anyfurther components. The hydrogenation process of the present inventioncan thus, in a further embodiment, include an upstream step comprisingthe preparation of the starting material to be hydrogenated by partialoxidation of a suitable hydrocarbon and the separation of the startingmaterial to be hydrogenated from the product stream obtained in thisway.

In particular, this starting material to be hydrogenated is MA.Preference is given to using MA which originates from the partialoxidation of hydrocarbons. Suitable hydrocarbon streams are benzene,C₄-olefins (e.g. n-butenes, C₄ raffinate streams) or n-butane.Particular preference is given to using n-butane since it represents aninexpensive, economical starting material. Processes for the partialoxidation of n-butane are described, for example, inUllmann'sEncyclopedia of Industrial Chemistry, 6th Edition, ElectronicRelease, Maleic and Fumarics Acids—Maleic Anhydride.

The reaction product obtained in this way is then taken up in a suitableorganic solvent or solvent mixture which has a boiling point atatmospheric pressure which is at least 30° C. above that of MA.

This solvent (absorption medium) is brought to a temperature in therange from 20 to 160° C., preferably from 30 to 80° C. The gas streamcomprising maleic anhydride from the partial oxidation can be broughtinto contact with the solvent in many ways: (i) passing the gas streaminto the solvent (e.g. via gas introduction nozzles or sparging rings),(ii) spraying the solvent into the gas stream and (iii) countercurrentcontact between the upflowing gas stream and the downflowing solvent ina tray column or packed column. In all three variants, the gasabsorption apparatuses known to those skilled in the art can be used.When choosing the solvent to be used, care has to be taken to ensurethat it does not react with the starting material, for example the MAwhich is preferably used. Suitable solvents are tricresyl phosphate,dibutyl maleate, high molecular weight waxes, aromatic hydrocarbonshaving a molecular weight of from 150 to 400 and a boiling point above140° C., for example dibenzylbenzene; dialkyl phthalates havingC₁-C₈-alkyl groups, for example dimethyl phthalate, diethyl phthalate,dibutyl phthalate, di-npropyl phthalate and diisopropyl phthalate;mono-, di-, tri- und tetraesters of cyclohexanedi-, tri- und tetraacids,the esters being alkyl-, cycloalkyl-, hydroxyund alkoxyalkylestershaving 1 to 30, preferably 2 to 20, in particular 3 to 18 carbon atomsand—in the case of non-cyclic groups—being linear or branched;non-limiting examples comprise: dialkyl cyclohexane-1,4-dicarboxylateswith identical alcohol groups, dialkyl cyclohexane1,3-dicarboxylateswith identical alcohol groups, dialkyl cyclohexane-1,2-dicarboxylateswith identical alcohol groups, mixed esters ofcyclohexane-1,2-dicarboxylic acid with C1 to C13-alcohols, mixed estersof cyclohexane-1,3-dicarboxylic acid with C1 to C13-alcohols, mixedesters of cyclohexane-1,4-dicarboxylic acid with C1 to C13-alcohols,alkyl esters of cyclohexane-1,2,4-tricarboxylic acid, alkyl esters ofcyclohexane-1,3,5-tricarboxylic acid, alkyl esters ofcyclohexane-1,2,3-tricarboxylic acid, alkyl esters ofcyclohexane-1,2,4,5-tetracarboxylic acid; mono-, di-, tri- undtetraesters of cyclohexenedi-, tri- und tetraacids, the esters beingalkyl-, cycloalkyl-, hydroxy- und alkoxyalkylesters having 1 to 30,preferably 2 to 20, in particular 3 to 18 carbon atoms and—in the caseof non-cyclic groups—being linear or branched; non-limiting examplescomprise: dialkyl cyclohexene-1,4-dicarboxylates with identical alcoholgroups, dialkyl cyclohexene-1,3-dicarboxylates with identical alcoholgroups, dialkyl cyclohexene-1,2-dicarboxylates with identical alcoholgroups, mixed esters of cyclohexene-1,2-dicarboxylic acid with C1 toC13-alcohols, mixed esters of cyclohexene-1,3-dicarboxylic acid with C1to C13-alcohols, mixed esters of cyclohexene-1,4-dicarboxylic acid withC1 to C13-alcohols, alkyl esters of cyclohexene-1,2,4-tricarboxylicacid, alkyl esters of cyclohexene-1,3,5-tricarboxylic acid, alkyl estersof cyclohexene-1,2,3-tricarboxylic acid, alkyl esters ofcyclohexene-1,2,4,5-tetracarboxylic acid; di-C₁-C₄-alkyl esters of otheraromatic and aliphatic dicarboxylic acids, for example dimethylnaphthalene-2,3-dicarboxylate, methyl esters of long-chain fatty acidshaving, for example, from 14 to 30 carbon atoms, high-boiling ethers,for example dimethyl ethers of polyethylene glycols, for exampletetraethylene glycol dimethyl ether.

The use of phthalates is preferred.

The solution resulting from treatment with the absorption mediumgenerally has an MA content of from about 5 to 400 grams per liter.

The waste gas stream remaining after treatment with the absorptionmedium comprises mainly the by-products of the preceding partialoxidation, e.g. water, carbon monoxide, carbon dioxide, unreactedbutanes, acetic acid and acrylic acid. The waste gas stream is virtuallyfree of MA.

The dissolved MA is subsequently stripped from the absorption medium.This is carried out using hydrogen at or at most 10% above the pressureof the subsequent hydrogenation or alternatively under reduced pressurewith subsequent condensation of remaining MA. In the stripping column, atemperature profile determined by the boiling points of MA at the top ofthe column and the virtually MA-free absorption medium at the bottom atthe prevailing column pressure and the chosen dilution with carrier gas(in the first case with hydrogen) is observed. In the case of directstripping with hydrogen, a temperature at the top of 130° C. and apressure of 5 bar are employed.

To prevent losses of solvent, rectification internals can be locatedabove the feed point for the crude MA stream. The virtually MA-freeabsorption medium taken off at the bottom is returned to the absorptionzone. In the case of direct stripping with hydrogen, a stream of gasvirtually saturated with MA is taken off at the top of the column at180° C. and a pressure of 5 bar. The H₂/MA ratio is from about 20 to400. Otherwise, the condensed MA is pumped to a vaporizer and vaporizedthere into the circulating gas stream.

The MA/hydrogen stream further comprises by-products formed in thepartial oxidation of n-butane, butenes or benzene by means ofoxygen-containing gases and also absorption medium which has not beenseparated off. These additional components are primarily acetic acid andacrylic acid as by-products, water, maleic acid and the dialkylphthalates which are preferably used as absorption media. The MAcontains from 0.01 to 1% by weight, preferably from 0.1 to 0.8% byweight, of acetic acid and from 0.01 to 1% by weight, preferably from0.1 to 0.8% by weight, of acrylic acid, based on MA. In thehydrogenation step, acetic acid and acrylic acid are completely orpartly hydrogenated to form ethanol or propanol. The maleic acid contentis from 0.01 to 1% by weight, in particular from 0.05 to 0.3% by weight,based on MA.

If dialkyl phthalates are used as absorption media, the amounts of thempresent in the MA depend strongly on correct operation of the strippingcolumn, in particular the enrichment section. Phthalate contents of 1.0%by weight, in particular 0.5% by weight, should not be exceeded undersuitable operating conditions, since otherwise the consumption ofabsorption media becomes too high.

The hydrogen/maleic anhydride stream obtained in this way is then passedto the hydrogenation zone and hydrogenated as described above. Comparedto the use of substantially prepurified, for example by distillation,MA, the catalyst activity and operating life is virtually unchanged. Theprocess of the present invention makes it possible to obtain THF yieldsof about 90%, in favorable cases about 95%. A high product selectivityis achieved at the same time. GBL is usually formed in amounts of lessthan 5%.

The process of the present invention is illustrated by the examplesbelow.

EXAMPLES Example 1

a) Catalyst Production

1.5 l of water are placed in a heatable precipitation vessel providedwith a stirrer and are heated to 80° C. Over a period of one hour, ametal salt solution comprising 731 g of Cu(NO₃)₂*2.5 H₂O and 1840 g ofAl(NO₃)₃*9 H₂O in 2000 ml of water and at the same time a 20% strengthby weight sodium carbonate solution are metered into this precipitationvessel while stirring until a pH of 8 is reached in the precipitationvessel. The mixture is stirred at this pH for another 15 minutes. Thetotal consumption of sodium carbonate solution is 5.6 kg. The suspensionformed is filtered and the solid is washed with water until the washingsno longer contain nitrate (<25 ppm). The filter cake is firstly dried at120° C. and subsequently calcined at 600° C. The catalyst produced inthis way comprises 50% by weight of CuO and 50% by weight of Al₂O₃. 400g of this catalyst powder are comminuted to a particle size of <1 mm,admixed with 12 g of graphite powder, intimately mixed and pressed toform pellets having a diameter of 3 mm and a height of 3 mm.

b) Catalyst Activation

Before the commencement of the reaction, the catalyst is subjected to atreatment with hydrogen in the hydrogenation apparatus. For thispurpose, the reactor is heated to 180° C. and the catalyst is activatedat atmospheric pressure for the time indicated in Table 1 using themixture of hydrogen and nitrogen specified in each case.

TABLE 1 Time (minutes) Hydrogen (Standard l/h) Nitrogen (Standard l/h)120 10 550 30 25 400 15 60 100 180 60 0

c) Hydrogenation Apparatus

The pressure apparatus used for the hydrogenation comprises a vaporizer,a reactor, a condenser with feed line for quench, a feed line forhydrogen, a waste gas line and a circulating gas blower. The pressure inthe apparatus is kept constant.

The molten MA is pumped from the top into the preheated (245° C.)vaporizer and vaporized. A mixture of fresh hydrogen and circulating gasis likewise introduced into the vaporizer from the top. Hydrogen and MAthus flow from the bottom to the heated reactor. The contents of thereactor comprise a mixture of glass rings and catalyst. After thehydrogenation, the THF formed together with water, other reactionproducts and hydrogen leaves the reactor and is condensed in thecondenser by quenching. Part of the circulating gas is bled off beforethe remainder, mixed with fresh hydrogen, flows back into the vaporizer.

The condensed liquid reaction mixture, the waste gas and the circulatinggas are analyzed quantitatively by gas chromatography.

Example 1d

d) Hydrogenation of Maleic Anhydride Prepared from n-butane

The reactor of the hydrogenation apparatus described in Example 1b ischarged with 220 ml of the catalyst produced in Example 1a and 130 ml ofglass rings. Activation was carried out as described in Example 1b.

The starting material used is maleic anhydride prepared from n-butaneand containing 500 ppm of acrylic acid, 1500 ppm of acetic acid and 100ppm of dibutyl phthalate. The reaction is carried out for 1000 hours.During the whole of this time, no deactivation of the catalyst, i.e. nodecrease in the maleic anhydride conversion and/or the tetrahydrofuranyield, is observed. Butanediol is not detected by gas chromatography.Table 2 summarizes the reaction parameters for the hydrogenation and theresults.

After fractional distillation of the hydrogenation products,tetrahydrofuran is isolated in a purity of 99.96%. This THF meets thespecification for use as starting material for poly THF.

Example 1e

In place of the catalyst comprising 50% by weight of CuO and 50% byweight of Al₂O₃, a catalyst having the composition 40% by weight of CuO,40% by weight of ZnO and 20% by weight of Al₂O₃ is used. This too is,after activation, operated for 1000 hours using the maleic anhydridedescribed in Example 1d without deactivation being observed. Butanediolis not detected by gas chromatography. Table 2 summarizes the reactionparameters for the hydrogenation and the results. After fractionaldistillation of the hydrogenation products, tetrahydrofuran is isolatedin a purity of 99.94%.

TABLE 2 WHSV Molar Temp. Pressure GHSV catalyst ratio Conversion³⁾ Mol %Example (° C.) (bar) (h⁻¹) (kg/lh) H₂:MA (%) THF GBL n-BuOH n-Butane1d¹⁾ 255-258 5.1 2000 0.1 85:1 100 87-88 4-5 6 0.1 1e²⁾ 2000 0.1 87 3 50.1 n-BuOH = n-butanol ¹⁾Catalyst: 50% by weight CuO, 50% by weightAl₂O₃ ²⁾Catalyst: 40% by weight CuO, 40% by weight ZnO, 20% by weightAl₂O₃ ³⁾Total conversion of MA and succinic anhydride (SA)

The results of Examples 1d and 1e show that an unchanged high catalystactivity together with unchanged tetrahydrofuran yields are achievedover long reaction times in the continuous hydrogenation of maleicanhydride in which acrylic acid, acetic acid and dibutyl phthalate arepresent over Cu/Al₂O₃ and Cu/ZnO/Al₂O₃ catalysts. In addition, thehydrogenation products can be worked up by distillation to givetetrahydrofuran of high purity which meets the required tetrahydrofuranspecification.

Example 2

A catalyst comprising 60% of copper oxide and 40% of aluminum oxide andproduced in a manner analogous to Example 1a is installed in theabove-described hydrogenation apparatus and pretreated with hydrogen asdescribed in Example 1b. The starting material used is maleic anhydrideprepared from n-butane and containing 5000 ppm of acrylic acid, 1500 ppmof acetic acid and 100 ppm of dibutyl phthalate. The reaction is carriedout at a pressure of 5 bar; all other reaction parameters and resultsare shown in Table 3. Examples 2a, 2b and 2d are comparative experimentscarried out under reaction conditions outside the parameters employedaccording to the present invention.

TABLE 3 WHSV Molar Temp. Pressure GHSV catalyst ratio Mol % Example (°C.) (bar) (h⁻¹) (kg/lh) H₂:MA THF GBL n-BuOH n-Butane 2a¹⁾ 235 5 50000.04 550:1 86.0 2.0 9.0 1.0 2b¹⁾ 2700 0.12 100:1 58.6 36.7 2.3 0.3 2c255 2700 0.13  90:1 89.4 2.6 7.0 1.0 2d¹⁾ 285 2700 0.13  90:1 65.3 —12.5 19.6 ¹⁾Comparative examples

The result of Example 2c shows that high tetrahydrofuran yields areachieved at 255° C. If, as in Example 2b, the temperature is reduced to235° C., the butyrolactone yield increases to 36.7%. Only by loweringthe WHSV over the catalyst to a third and increasing the hydrogen/maleicanhydride ratio (Example 2a) is a tetrahydrofuran yield approaching thatof Example 2c obtained. If (Example 2d) the hydrogenation temperature isincreased to 285° C., the tetrahydrofuran yield drops to 65.3%.

Example 3

Example 2c is repeated at a GHSV of 1800 h⁻¹ under otherwise identicalconditions using a catalyst comprising 60% of CuO and 40% of Al₂O₃. Thetetrahydrofuran yield is 93%. Butyrolactone is not observed.

Example 4 (Comparative Example)

Example 3 of JP 2-233 631 is repeated: for this purpose, the Cu/Al₂O₃catalyst described in Example 3 is produced by the method indicatedthere. The pure MA/GBL mixture is then hydrogenated under the conditionsindicated. This gives a THF yield of 90% and a GBL yield of 7%, also1.9% of n-butanol.

On changing from MA/GBL (molar ratio=1:3) to pure MA (molar amount ofpure MA corresponds to the sum of the molar amounts of MA+GBL, WHSV overthe catalyst=0.03 kg of MA/l of catalyst x hour) as feed, the THF yielddrops to 12% of THF and 59% of GBL after 10 hours. In addition, 28% ofmaleic acid and succinic acid (based on MA used) are formed (Table 4).Large amounts of dicarboxylic acid mixture were deposited in the outletsection of the hydrogenation apparatus and in the circulating gassystem.

TABLE 4 MA conver- Example sion (%) (Mol %)²⁾ 4¹⁾ About 95 THF GBL BDO³⁾n-BuOH n-Butane 12 59 — 0.6 — ¹⁾Comparative example ²⁾Based on MA used³⁾1,4-Butanediol

The results show that only very low THF yields are achieved under theconditions of Example 3 of JP 2-233 631 when using MA in place of MA/GBLmixtures. Despite low WHSVs over the catalyst, blockages caused bysolids (dicarboxylic acid mixtures) occur in the hydrogenation plant andthese make it impossible to carry out the MA hydrogenation on anindustrial scale.

We claim:
 1. A process for preparing unsubstituted or alkyl-substitutedTHF by catalytic hydrogenation in the gas phase of C₄-dicarboxylic acidsor their derivatives using a catalyst comprising <80% by weight of CuOand >20% by weight of an oxidic support having acid centers, at a hotspot temperature of from 240 to 310° C. and a WHSV over the catalyst offrom 0.01 to 1.0 kg of starting material/1 of catalyst x hour.
 2. Aprocess as claimed in claim 1, wherein the catalyst comprises 10 to 65%by weight of CuO an oxidic support.
 3. A process as claimed in claim 1,wherein the catalyst comprises 35 to 90% by weight of CuO anoxidicsupport.
 4. A process as claimed in claim 1, wherein the oxidic supportis Al₂O₃ or a combination of Al₂O₃/ZnO in a weight ratio of from 20:1 to1:20.
 5. A process as claimed in claim 1, carried out at pressures offrom 1 to 30 bar.
 6. A process as claimed in claim 1, wherein the molarratio of hydrogen/starting material is from 20 to
 400. 7. A process asclaimed in claim 1, wherein the GHSV is from 100 to 10,000 Standardm³/m³h.
 8. A process as claimed in claim 1, wherein the inlettemperature is from >220 to 300° C. and is from about 5 to 15° C. belowthe hot spot temperature.
 9. A process as claimed in claim 1, whereinthe hot spot is located in the first half of the reactor.
 10. A processas claimed in claim 1, wherein one or more further metals or compoundsthereof, from the group consisting of the elements of groups 1 to 14 ofthe Periodic Table of the Elements are present in the catalyst.
 11. Aprocess as claimed in claim 10, wherein a substance selected from thegroup consisting of ZrO₂, TiO₂, SiO₂ and MgO is present in the catalyst.12. A process as claimed in claim 1, wherein the catalyst is activatedby reduction before or after installation in the reactor and before usein the hydrogeneration reaction.
 13. A process as claimed in claim 1,wherein the catalyst further comprises an auxiliary in an amount of <10%by weight, selected from graphite, stearic acid, silica gel and copperpowder.
 14. A process as claimed in claim 1, wherein the shaped catalystbody has a pore volume of ≧0.01 ml/g for pore diameter of >50 nm.
 15. Aprocess as claimed in claim 1, wherein the ratio of macropores having adiameter of >50 nm to the total pore volume for pores having a diameterof >4 nm in the shaped catalyst body is >10%.
 16. A process as claimedin claim 1, wherein a fixed-bed reactor, a shaft reactor, afluidized-bed reactor or a reactor with internal heat removal is used.17. A process as claimed in claim 1, wherein maleic anhydride is used asstarting material for the reaction.
 18. A process as claimed in claim 1,wherein maleic anhydride prepared by oxidation of benzene, C₄-olefins orn-butane, where the crude maleic anhydride obtained by oxidation hasbeen extracted from the crude product mixture using a solvent and hassubsequently been stripped from this solvent by means of hydrogen, isused.
 19. A process as claimed in claim 1, wherein the absorption mediumis selected from the group consisting of tricresyl phosphate, dibutylmaleate, high molecular weight waxes, aromatic hydrocarbons having amolecular weight in the range from 150 to 400 and a boiling point above140° C., di-C₁-C₄-alkyl esters of aromatic and aliphatic dicarboxylicacids, methyl esters of long-chain fatty acids having from 14 to 30carbon atoms, high-boiling ethers, dimethyl ethers of polyethyleneglycols, and dialkyl phthalates having C₁-C₈-alkyl groups.
 20. A processas claimed in claim 1, wherein the maleic anhydride is stripped from theabsorption medium under reduced pressure or at pressures whichcorrespond to the pressure in the hydrogenation or are at most 10% abovethis pressure.